Configurations and methods of acid gas removal

ABSTRACT

A plant includes an absorber ( 105 ) that receives a feed gas ( 10 ) at a pressure of at least 400 psig and comprising at least 5 mol % carbon dioxide, wherein the absorber ( 105 ) is operated at an isothermal or decreasing top-to-bottom thermal gradient, and wherein the absorber ( 105 ) employs a physical solvent to at least partially remove an acid gas from the feed gas ( 10 ). Such configuration advantageously provide cooling ( 108 ) by expansion of the rich solvent ( 21 ) generated in the absorber ( 105 ), wherein both a semi-rich solvent ( 13 ) generated and recycled to the absorber ( 105 ) and the feed gas ( 10 ) are cooled by expansion of the rich solvent ( 21 ).

FIELD OF THE INVENTION

The field of the invention is removal of acid gases from a feed gas, andparticularly relates to acid gas removal from high carbon dioxidecontent feed gas.

BACKGROUND OF THE INVENTION

Acid gas removal from various gas streams, and especially removal ofcarbon dioxide from natural gas streams has become an increasinglyimportant process as the acid gas content of various gas sources isrelatively high, or increases over time. For example, there arerelatively large natural gas resources (e.g., Alaska, Continental NorthAmerica, Norway, Southeast Asia, or Gulf of Mexico) that contain highconcentrations of carbon dioxide ranging from 20% to 75%. Moreover,where enhanced oil recovery (EOR) is employed, the carbon dioxideconcentration in natural gas will increase over time to significantconcentrations that will typically require gas processing to remove atleast part of the carbon dioxide.

Currently more than half of the natural gas produced in the U.S. istreated to meet pipeline specification with minimal processing, and suchprocessing frequently includes glycol dehydration and hydrocarbonremoval. Untreated gas with high carbon dioxide content is usually leftin the ground, mostly due to economical and/or technical considerations.

Among other difficulties, removal of impurities (primarily water,hydrogen sulfide, and/or carbon dioxide) is generally required totransport the treated natural gas through pipelines, which significantlyincreases production costs. Furthermore, many known acid gas removalprocesses also remove a portion of the methane and other hydrocarbons.(Losses of less than about 2% of hydrocarbons are normally acceptable,losses of 5-10% may be acceptable if the value of the product gas ishigh or offset by other advantages, while losses above 10% are normallyunacceptable). Still further, the removed carbon dioxide must typicallybe recompressed back to the high pressure formation to reduce itsenvironmental impact and for enhanced oil recovery, which is energyintensive and therefore economically unattractive.

To overcome at least some of the disadvantages associated with acid gasremoval, numerous processes were developed and may be categorized intovarious categories, wherein the choice of the appropriate gas treatmentwill predominantly depend on the gas composition, the size and locationof the plant, and other variables.

For example, in one category one or more membranes are used tophysically separate the acid gas from a gaseous feed stream, wherein atypical membrane system includes a pre-treatment skid and a series ofmembrane modules. Membrane systems are often highly adaptable toaccommodate treatment of various gas volumes and product-gasspecifications. Furthermore, membrane systems are relatively compact andare generally free of moving parts, therefore rendering membrane systemsan especially viable option for offshore gas treatment. However, all oralmost all single stage membrane separators are relatively non-selectiveand therefore produce a carbon dioxide permeate stream with a relativelyhigh methane and hydrocarbon content (which is either vented,incinerated or used as a low BTU fuel gas). Consequently, the highmethane and hydrocarbon losses tend to render the use of this processundesirable and uneconomical. To reduce such losses, multiple stages ofmembrane separators with inter-stage recompression may be used. However,such systems tend to be energy intensive and costly.

In another category, a chemical solvent is employed that reacts with theacid gas to form a (typically non-covalent) complex with the acid gas.In processes involving a chemical reaction between the acid gas and thesolvent, the crude gases are typically scrubbed with an alkaline saltsolution of a weak inorganic acid (e.g., U.S. Pat. No. 3,563,695 toBenson), or with an alkaline solution of organic acids or bases (e.g.,U.S. Pat. No. 2,177,068 to Hutchinson). One particular advantage of achemical solvent system is that such systems typically absorb methane toa relatively low degree. Furthermore, chemical solvent systems oftenproduce a product gas with a very low acid gas content.

However, while use of chemical solvent systems may be advantageous in atleast some respects (see above), substantial difficulties are frequentlyinherent. For example, once the chemical solvent is spent, the acid gasis flashed off and the solvent is regenerated using heat, which may addsubstantial cost to the acid gas removal. Furthermore, the mechanicalequipment in a gas treatment plant using a chemical solvent is oftenprone to failure from either corrosion or foaming problems. Stillfurther, chemical solvent systems typically include columns, heaters,air coolers, pumps, etc., all of which require frequent quality checksand maintenance, making operational reliability probably the weakestfeature of such systems. Yet another disadvantage of chemical solventsystems is that the product gas and carbon dioxide streams musttypically be further dried to meet pipeline specifications. Moreover,the quantity of chemical solvent required to absorb increasing amountsof acid gases generally increases proportionally with acid gas quantity,thus making the use of chemical solvents problematic where the acid gascontent increases over time in the feed gas.

In a still further category, a physical solvent is employed for removalof acid gas from a feed gas, which is particularly advantageous fortreating gas with a high acid gas partial pressure as the potentialtreating capacity of the physical solvent increases with the acid gaspartial pressure (Henry's law). Using physical solvents, absorption of aparticular acid gas predominantly depends upon the particular solventemployed, and is further dependent on pressure and temperature of thesolvent. For example, methanol may be employed as a low-boiling organicphysical solvent, as exemplified in U.S. Pat. No. 2,863,527 to Herbertet al. However, the refrigerant cooling requirement to maintain thesolvent at cryogenic temperatures is relatively high, and the processoften exhibits greater than desired methane and ethane absorption,thereby necessitating large energy input for recompression and recovery.

Alternatively, physical solvents may be operated at ambient or slightlybelow ambient temperatures, including propylene carbonates as describedin U.S. Pat. No. 2,926,751 to Kohl et al., and those usingN-methylpyrrolidone or glycol ethers as described in U.S. Pat. No.3,505,784 to Hochgesand et al. In further known methods, physicalsolvents may also include ethers of polyglycols, and specificallydimethoxytetraethylene glycol as shown in U.S. Pat. No. 2,649,166 toPorter et al., or N-substituted morpholine as described in U.S. Pat. No.3,773,896 to Preusser et al. While use of physical solvents avoids atleast some of the problems associated with chemical solvents ormembranes, various new difficulties arise. Among other things, mostknown solvent processes lack an efficient heat exchange integrationconfiguration, and often require significant refrigeration and/or highsolvent circulation, and sometimes require heat for solventregeneration. In most or almost most of the known physical solventprocesses, co-absorption of methane and hydrocarbons can be relativelyhigh due to the high solvent circulation.

Furthermore, where relatively low carbon dioxide content in the productgas is required, various physical solvent processes require steam orexternal heat for solvent regeneration. A typical physical solventprocess is exemplified in Prior Art FIG. 1, which is conceptuallyrelatively simple and employs use of a cold lean solvent to removecarbon dioxide from the feed gas. The solvent is regenerated bysuccessive flashing to lower pressures and the flashed solvent is thenpumped to the absorber, wherein the solvent is cooled using externalrefrigeration (either in the rich solvent or the lean solvent circuit).In most instances, a steam or fuel fired heater is required for solventregeneration.

In such processes, as carbon dioxide is absorbed by the solvent, theheat of solution of carbon dioxide increases the solvent temperatureresulting in a top-to-bottom increasing temperature profile across theabsorber. Consequently, one limitation of physical absorption lies inthe relatively high absorber bottom temperature, which limits carbondioxide absorption capacity of the solvent. To overcome the problemsassociated with limited absorption capacity, the solvent circulationrate may be increased. However, increase in solvent circulationsignificantly increases refrigeration costs and energy consumption forpumping the solvent. Worse yet, high solvent circulation of knownsolvent processes will lead to increased loss of methane andhydrocarbons (due to co-absorption). Yet another undesirable aspect ofknown physical solvent processes is problematic heat and mass transferdue to the cold lean solvent temperature entering the top of theabsorber: While a relatively cold lean solvent is required to reducesolvent circulation in known processes, further reduction of the leansolvent temperature becomes undesirable as the solvent's surface tensionand viscosity increase, eventually leading to hydraulic problems.

Moreover, in all or almost all of the known acid gas removal processesusing solvents the acid gas is removed in the regenerator at low orsubstantially atmospheric pressure. Consequently, and especially wherethe carbon dioxide is later used for EOR, the isolated carbon dioxidemust be compressed to substantial pressures, which further increasesprocess costs. Thus, although various configurations and methods areknown to remove acid gases from a feed gas, all or almost all of themsuffer from one or more disadvantages. Therefore, there is still a needto provide methods and configurations for improved acid gas removal.

SUMMARY OF THE INVENTION

The present invention is directed towards methods and configurations ofa plant comprising an absorber that receives a feed gas at a pressure ofat least 400 psig with at least 5 mol % carbon dioxide, wherein theabsorber is operated at an isothermal or decreasing top-to-bottomthermal gradient, and wherein the absorber employs a physical solvent toat least partially remove an acid gas from the feed gas.

The absorber in preferred plants produces a semi-rich solvent and a richsolvent, and wherein the semi-rich solvent is cooled by at leastpartially expanded rich solvent. It is further contemplated thatpreferred absorbers produce a rich solvent that is expanded in at leasttwo steps, wherein expansion in one step produces work, and whereinexpansion in another step provides refrigeration for at least one of asemi-rich solvent produced by the absorber and a carbon dioxide product.In further aspects of preferred absorbers, the absorber produces a richsolvent that is expanded in at least three steps, wherein expansion inthe at least three steps produces at least three recycle streams,respectively, and wherein the at least three recycle streams (which mayfurther be compressed to form a compressed recycle stream, and whereinfurther refrigeration may be provided by Joule-Thomson cooling ofcompressed recycle stream) are fed into the absorber. It is furthercontemplated that the absorber is operated at a bottom temperature ofabout −25° F. to about 45° F. and produces a rich solvent that isexpanded to provide refrigeration for a carbon dioxide product.

Therefore, contemplated absorbers may receive a natural gas comprisingat least 5 mol % acid gas and having a pressure of at least 400 psig andinclude a physical solvent that absorbs at least a portion of the acidgas in the absorber to form a semi-rich solvent, wherein a cooler isfluidly coupled to the absorber that receives and cools the semi-richsolvent and provides the cooled semi-rich solvent back to the absorber,wherein the cooled semi-rich solvent further absorbs at least anotherportion of the acid gas to form a rich solvent, and wherein the naturalgas and the semi-rich solvent are cooled at least in part by expansionof the rich solvent.

Various objects, features, aspects and advantages of the presentinvention will become more apparent from the following detaileddescription of preferred embodiments of the invention, along with theaccompanying drawing.

BRIEF DESCRIPTION OF THE DRAWING

Prior art FIG. 1 is schematic depicting an exemplary known configurationfor acid gas removal using a physical solvent.

FIG. 2 is one exemplary schematic depicting a plant configuration foracid gas removal according to the inventive subject matter.

FIG. 3 is another exemplary schematic depicting a plant configurationfor acid gas removal for EOR with additional carbon dioxide liquidproduction using internally produced refrigeration.

FIG. 4 is a further exemplary schematic depicting a plant configurationfor acid gas removal for EOR configured with additional membraneseparation upstream and carbon dioxide liquid production usinginternally produced refrigeration.

DETAILED DESCRIPTION

The inventors have discovered that acid gases, and particularly carbondioxide, may be removed from a feed gas comprising at least 5 mol %carbon dioxide using configurations and methods in which an absorberreceives a feed gas at a pressure of at least 400 psig, wherein theabsorber is operated at an isothermal or decreasing top-to-bottomthermal gradient, and wherein the absorber employs a physical solvent toat least partially remove an acid gas from the feed gas.

As used herein, the term “isothermal gradient” means that thetemperature of the physical solvent in an upper portion of the absorberis substantially identical (i.e., absolute deviation of temperature nomore than 10° F.) with the temperature of the physical solvent in amiddle and lower portion of the absorber. Similarly, the term“decreasing top-to-bottom thermal gradient” as used herein means thatthe temperature of the physical solvent in an upper portion of theabsorber is higher than the temperature of the physical solvent in amiddle and/or lower portion of the absorber.

As further used herein, and with respect to a column or absorber, theterms “upper” and “lower” should be understood as relative to eachother. For example, withdrawal or addition of a stream from an “upper”portion of a column or absorber means that the withdrawal or addition isat a higher position (relative to the ground when the column or absorberis in operation) than a stream withdrawn from a “lower” region thereof.Viewed from another perspective, the term “upper” may thus refer to theupper half of a column or absorber, whereas the term “lower” may referto the lower half of a column or absorber. Similarly, where the term“middle” is used, it is to be understood that a “middle” portion of thecolumn or absorber is intermediate to an “upper” portion and a “lower”portion. However, where “upper”, “middle”, and “lower” are used to referto a column or absorber, it should not be understood that such column isstrictly divided into thirds by these terms.

As still further used herein, the term “about” when used in conjunctionwith numeric values refers to an absolute deviation of less or equalthan 10% of the numeric value, unless otherwise stated. Therefore, forexample, the term “about 10 mol %” includes a range from 9 mol %(inclusive) to 11 mol % (inclusive).

In a preferred configuration as depicted in FIG. 2, an exemplary plantcomprises a gas pretreatment unit 101 that may include (1) one or moregas coolers for removing the bulk of the water content by cooling thegas to just above the gas hydrate temperature typically at 60° F.; (2)heavy hydrocarbon removal units for C₆+ components in a feed gas; (3) agas dehydration unit, preferably a molecular sieve unit or a glycol unit(Drizo) to produce a very low water dew-point feed gas. Water and heavyhydrocarbons are removed in the pretreatment unit 101 from the feed gasstream 1 as water and heavy hydrocarbon stream 3 to form treated feedgas stream 2.

It is particularly preferred that treated feed gas stream 2 is furthercooled to typically 10° F. to 40° F. in a heat exchanger 103 usingabsorber overhead stream 11 as a refrigerant to form cooled treated feedgas stream 7, which is mixed with combined recycle stream 8 to formstream 9 that is further cooled in heat exchanger 104. In thisconfiguration, heat exchanger 104 uses refrigeration provided byatmospheric depressurized rich solvent stream 28 and further coolsstream 9 to typically −15° F. to −45° F. thereby forming cooled stream10. The so cooled stream 10 enters the absorber 105 at a lower portionof the absorber. It should be particularly appreciated that cooling ofthe treated feed gas stream to a relatively low temperature (e.g., about−15° F. to about −45° F.) will maintain the absorber bottom temperatureat a particularly low level (e.g., about 0° F. to about −40° F.), whichis used to maximize the carbon dioxide loading of the rich solvent, andthereby to minimize solvent circulation and methane and hydrocarbonslosses.

It is further preferred that a side cooler 108 is employed to controland/or maintain the temperature of the lower section of the absorber 105at a predetermined absorption temperature. In such configurations, thesemi-rich solvent stream 13 (generated by absorption of acid gas in anupper portion of the absorber) is pumped by the side cooler pump 106(stream 14) and is cooled in heat exchanger 108 using flashed richsolvent stream 21 from hydraulic turbine 111 as refrigerant. The socooled semi-lean solvent stream 15, at typically −10° F. to −40° F., isreturned to the lower section of the absorber 105.

It is especially preferred that the refrigerant for the side cooler 108is provided by the flashed rich solvent stream 21 (depressurized richsolvent stream) from hydraulic turbine 111. However, it should berecognized that cooling may also be provided by various otherrefrigerant, and suitable refrigerants may be internal (produced withinthe plant) or external. For example, the refrigerant for side cooler 108may also be provided by flashing of carbon dioxide and/or via anexpansion turbine. Alternatively, refrigeration may be provided and/orsupplemented by JT cooling created from the recycle gas cooler 125 andJT 140, or by an external source at exchanger 102 with an externalrefrigerant 37, particularly when the feed gas pressure is low.

It should be especially appreciated that when a heavier gas isprocessed, a hydrocarbon liquid stream 150 is formed at the discharge ofthe recycle gas cooler 125. Recovery of such liquid products will add tothe economical benefit of this process while reducing the gas recycle.

Thus, suitable side coolers may be advantageously configured to maintainan optimum absorption temperature for effective absorption of the acidgas. Consequently, it should be recognized that in such configurationsthe middle portion of the absorber is preferably operated at a lowertemperature than the upper portion of the absorber, which isparticularly advantageous when the solvent is loaded with carbon dioxide(the solvent will typically exhibit lower viscosity and lower surfacetension).

The semi-rich solvent will then in the absorber further absorb carbondioxide from the feed gas, thereby forming rich solvent 16 that exitsthe absorber via first hydraulic turbine 107.

First hydraulic turbine 107 reduces the absorber bottoms pressure totypically about half of the feed gas pressure, thus cooling the richsolvent to about −5° F. to −35° F. to form a depressurized rich solventstream 17. It is generally contemplated that the hydraulic turbine is anenergy efficient device as it generates refrigeration cooling byexpansion and flashing of the carbon dioxide content while providingshaft work to drive the solvent circulation pump.

The rich solvent 17 is flashed to separator 110 which produces a firstflashed hydrocarbon vapor (first hydrocarbon recycle stream 19) that isrecovered to the absorber 105 via recycle compressor 124 and stream 8.The so flashed solvent stream 20 is further expanded in a secondhydraulic turbine 111 to a pressure reduced by half to form an expandedrich solvent stream 21 (typically at −20° F. to −40° F.), which is usedto cool the semi-rich solvent stream 14 in heat exchanger 108. Theheated rich solvent 22 from heat exchanger 108, typically at 10° F. to−10° F., is separated in separator 112, which produces a second flashedhydrocarbon vapor (second hydrocarbon recycle stream 23) to be recycledvia recycle compressor 124. The flashed liquid stream 24 from separator112 is further let down in pressure in an expansion JT valve 113 toreduce pressure typically by half, thereby chilling the rich solvent to5° F. to −15° F. The so flashed solvent 25 is separated in separator 114which produces a third flashed hydrocarbon vapor (third hydrocarbonrecycle stream 26) to be recycled via recycle compressor 124. The powergenerated from the first and second hydraulic turbines 107 and 111 canbe used to provide part of the power requirement of the lean solventpump 119, vacuum pump 120, recycle compressor 124 or for powergeneration.

The flashed liquid 27 from separator 114 is let down in pressure in anexpansion JT valve 115 to above atmospheric pressure, thereby furtherchilling the rich solvent to −20° F. to −45° that is then used forchilling the feed gas in heat exchanger 104. The heated rich solvent 29from heat exchanger 104, typically at 0° F. to −40° F., is thenseparated in separator 116 at atmospheric pressure to produce a flashedcarbon dioxide stream 30 that can vented or used for enhanced oilrecovery. To further enhance solvent regeneration efficiency, theatmospheric flashed solvent 31 is expanded via JT valve 117 to vacuumpressures (typically 1 to 10 psia) in stream 32, which is separated invacuum separator 118 to form an ultra lean solvent stream 34 and aflashed carbon dioxide vapor stream 33. The ultra lean solvent 34 ispumped by lean solvent pump 119 to the absorber pressure for carbondioxide absorption and delivered via compressed ultra lean solventstream 35. The carbon dioxide may then be compressed via vacuum pump 120to form compressed carbon dioxide stream 37.

It should be particularly appreciated that the so generated carbondioxide streams will contain over 95 mol % CO₂, which are suitable forenhanced oil recovery. If necessary, higher purity carbon dioxide streamcan be produced by increasing the temperatures and/or reducing thepressures of the flash separators. The product gas, to meet pipelinecarbon dioxide specification, contains typically 2% CO₂, which canfurther be reduced with the use of an ultra lean solvent formed byfurther reducing the vacuum pressure of the vacuum separator and withthe use of a stripping gas (via vacuum stripping).

Where enhanced oil recovery is particularly desirable, it iscontemplated that configurations according to the inventive subjectmatter may be modified as depicted in FIG. 3, in which like numeralsdepict like components as shown in FIG. 2. In plant configurationsaccording to FIG. 3, an additional heat exchanger 109 is employed tocool the carbon dioxide stream 41 using the depressurized rich solventstream 17 from the hydraulic turbine 107. In addition, the flashedcarbon dioxide vapor 33 is compressed in a vacuum pump 120 toatmospheric pressure, combined with stream 36 to form stream 38, andstill further compressed in compressor 121. The compressed carbondioxide stream 39 is cooled to its liquid state (in stream 43)successively by heat exchangers 122, 123, and 109. An optional trimcondenser 124 with external refrigeration (44) may be required tosupplement refrigeration duty required by carbon dioxide condensation.Carbon dioxide liquid 43 is pumped by pump 125 to stream 46 forre-injection for enhanced oil recovery, typically at 4000 psig.

Alternatively, and especially where the feed gas pressure is relativelyhigh (e.g., above 1000 psig), an upstream membrane separation unit maybe employed as depicted in FIG. 4, in which like numerals depict likecomponents as shown in FIG. 3. In such configurations, it is especiallypreferred that one or more membrane separators 102 perform as a bulkcarbon dioxide removal unit producing a non-permeate stream 5 withcarbon dioxide content typically 30% to 50% at a high pressure, and apermeate stream 4 with carbon dioxide content typically at 60% to 95%and at a permeate pressure that can be combined with the other carbondioxide stream for enhanced oil recovery. Of course, it should berecognized that a particular carbon dioxide concentration will at leastin part depend on the particular membrane separator used, and further onthe solvent unit and treating specifications of the product gas and thecarbon dioxide stream. Alternatively, or additionally, a portion of thepermeate stream may also be suitable for use as regeneration gas orstripping gas for the dehydration unit.

Thus, it should be especially recognized that the carbon dioxide contentin the feed gas will provide refrigeration for solvent chilling as wellas liquefaction duty of the carbon dioxide stream by the expansion ofthe rich solvent with hydraulic turbines and JT valves. It shouldfurther be appreciated that if additional refrigeration is required(e.g., at relatively low feed pressure), solvent cooling can be suppliedby JT cooling with the recycle gas compressor 124 compressing to ahigher pressure, cooled in heat exchanger 125 and letdown using JT valve140 to the absorber pressure.

With respect to suitable feed gases, it is contemplated that numerousnatural and synthetic feed gases are appropriate. However, particularlypreferred feed gases include natural gas, and especially natural gaswith a carbon dioxide that is at least about 5 mol %, more typically atleast 10 about mol %, and most typically at least 10 to 75 mol %.Therefore, especially suitable feed streams include natural gas feedstreams from oil and gas fields such as Alaska, Norway, Southeast Asiaand Gulf of Mexico. Similarly, the acid gas content (and especiallycarbon dioxide content) of suitable feed gases may vary and willpredominantly depend on the source of the feed gas. It is generallypreferred, however, that the acid gas content will be at least about 5mol %, more typically at least 10 about mol %, and most typically atleast 20 to 75 mol %. A typical feed gas composition is given in Table 1below:

TABLE 1 COMPONENT MOL % N₂ 0.88 CO₂ 19.14 H₂S 0.00 C₁ 72.69 C₂ 5.29 C₃1.40 IC₄ 0.22 NC₄ 0.26 IC₅ 0.02 NC₅ 0.01 C₆ 0.08

Furthermore, it should be recognized that the pressure of contemplatedfeed gases may vary considerably, and suitable pressures will rangebetween atmospheric pressure and several thousand psig. However, it isparticularly preferred that the feed gas has a pressure of at least 400psig, more typically at least 1000 psig, even more typically at least3000 psig, and most typically at least 5000 psig. Moreover, while it isgenerally contemplated that at least a portion of the feed gas pressureis due to the pressure of the gas contained in the well, it should alsobe recognized that where appropriate, the pressure may also be increasedusing one or more compressors.

In yet further aspects of the inventive subject matter, contemplated fedgases are preferably cooled before entering the absorber, and it isespecially preferred that the cooling of the feed gas will be at leastin part effected by the product gas (i.e., the absorber overhead stream)in one or more heat exchangers. With respect to the degree of cooling,it is generally contemplated that the feed gas may be cooled to varioustemperatures. However, it is especially preferred that the feed streamwill be cooled to a temperature just above the gas hydrate point. Thecooled feed gas stream may then be fed into a separator in which atleast a portion of the water contained in the feed gas is removed fromthe cooled feed stream to form a partially dehydrated feed gas.

The so formed partially dehydrated feed gas may then be further treatedto remove higher hydrocarbons (e.g., C₆ ⁺) and then still furtherdehydrated in a dehydration unit (all known gas dehydration units aresuitable for use). For example, further dehydration may be performedusing glycol or molecular sieves. Dehydration of the feed isparticularly advantageous because the absorption process can be run atsignificantly lower temperature. Moreover, the product gas and thecarbon dioxide are produced in a very dry state that eliminates anydownstream dehydration of the product gases.

In still further preferred aspects, and especially where the feed gaspressure and/or carbon dioxide content is relatively high, it iscontemplated that the dehydrated feed gas may be further separated inmembrane separators to produce a carbon dioxide rich permeate that canbe used for enhance oil recovery or regeneration gas of the dehydrationunit and a non-permeate for downstream solvent absorption. However, theuse of the membrane separators may not be required when carbon dioxidecontent in the feed gas is less than 50%, most typically at least 10% to45%. In especially preferred contemplated configurations using membraneseparators, the dried non-permeate is cooled in a first heat exchanger,wherein the cooling duty is provided by the product gas (i.e., theabsorber overhead stream), and in a second heat exchanger, wherein thecooling is further provided by the expanded rich solvent. Membraneseparation technology is attractive for this separation, becausetreatment can be accomplished using the high wellhead gas pressure asthe driving force for the separation. Conventional membrane separatorssuch as the cellulose acetate membranes can provide adequate selectivityfor carbon dioxide removal with minimal methane loss in the permeatestream.

Therefore, it should be particularly recognized that suitable absorberswill operate at relatively high pressure, and especially contemplatedhigh pressures are at least 500 psi, typically at least 1000 psi, evenmore typically at least 3000 psi, and most typically at least 5000 psi.Consequently, it should be recognized that contemplated absorbers mayoperate in a gas phase supercritical region. The term “operate in a gasphase supercritical region” as used herein refers to operation of theabsorber under conditions in which at least a portion of the feed gas,if not all of the feed gas, will be in a supercritical state.Furthermore, by operating the absorption process in the gas phasesupercritical region, hydrocarbon condensation is typically avoided,which currently presents a significant problem in heretofore knownprocesses. In yet further contemplated aspects, the type of absorberneed not be limited to a particular configuration, and all knownabsorber configurations are deemed suitable for use herein. However,particularly preferred contacting devices include a packed bed or trayconfigurations.

With respect to the solvent employed in contemplated absorbers, itshould be recognized that all physical solvents and mixtures thereof areappropriate. There are numerous physical solvents known in the art, andexemplary preferred physical solvents include propylene carbonate,tributyl phosphate, normal methylpyrrolidone, dimethyl ether ofpolyethylene glycol, and/or various polyethylene glycol dialkyl ethers.Alternatively, other solvents including enhanced tertiary amine (e.g.,piperazine) having similar behavior as physical solvent may be employed.

Consequently, the absorber will provide a product gas that is depletedof acid gases, and particularly depleted of carbon dioxide. Moreover, itshould be recognized that since the absorber receives a cooled anddehydrated feed gas, the product gas would typically conform to all oralmost all sales gas specifications and requirements for transportationthrough high-pressure pipelines. It should further be especiallyappreciated that the rich solvent formed in the absorber may leave theabsorber bottom at relatively high pressure (e.g., at least 500 psi,more typically between 1000 and 3000 psi), and may thus be utilized toprovide work (e.g., for generation of electrical energy) and/or coolingof various streams in the separation process.

In especially preferred configurations, the rich solvent is let down inpressure using a first hydraulic turbine to generate mechanical orelectric energy, and the depressurized rich solvent is then separated ina separator into a hydrocarbon-containing first recycle stream and afirst rich solvent, which is subsequently (optionally) employed as acoolant to refrigerate a carbon dioxide stream for the enhanced oilrecovery application (wherein the carbon dioxide is produced from thefeed gas). The hydrocarbon-containing first recycle stream is preferablyrecycled to the absorber, white the first rich solvent is furtherdepressurized using a second hydraulic turbine to further generatemechanical or electric energy. The so further depressurized rich solventstream is then employed as a refrigerant in a heat exchanger (preferablya side cooler of the absorber) that cools the semi-rich solvent in theabsorber to maintain a desirable absorber temperature. After passingthrough the heat exchanger, the further depressurized rich solventstream is then separated in a second separator into a second richsolvent and a second hydrocarbon-containing recycle stream that isrecycled to the absorber. From the second separator, the rich solventstream is further depressurized by a JT valve and then separated in athird separator into a third rich solvent and a thirdhydrocarbon-containing recycle stream that is recycled to the absorber.The third depressurized rich solvent is then further depressurized toatmospheric pressure, generating refrigeration that is to be used tocool the feed gas, maintaining the absorber at a desirable low bottomtemperature.

With the refrigeration provided by depressurizing the rich solvent,supplemental refrigeration is not required in most cases (particularlyin high feed pressure operation). If extra refrigeration is required, itmay be obtained internally by JT cooling created from the recycle gascooler and JT valve, or from an external source in an exchanger with arefrigerant. Furthermore, the particular heat exchanger sequence mayvary depending on the feed gas, solvent circulation and the carbondioxide liquefaction duty requirements. For example, the firstdepressurized rich solvent may be used to chill the feed gas instead ofthe carbon dioxide stream, and the second depressurized rich solvent maybe used for condensation of the carbon dioxide stream instead of theside cooler, and the third depressurized rich solvent cooler may be usedfor the side cooler instead of cooling the feed gas. Consequently, inpreferred configurations a lean solvent is formed at higher temperatureswith desirable thermal physical properties that enhance the hydrodynamicperformance of the absorption process, and a rich solvent at the lowestpossible temperature that maximizes carbon dioxide holding capacity ofthe solvent. Therefore, contemplated processes will result in lowersolvent circulation, lower methane and hydrocarbons losses, and lowerenergy consumption than currently known solvent based acid gas removalprocesses.

Flashing of the rich solvent may be performed in various configurations,and it is generally contemplated that all known configurations aresuitable for use herein. However, it is typically preferred that therich solvent (after providing work and/or cooling) is further let downin pressure to a pressure sufficient to release at least 80% (moretypically at least 90%, and most typically at least 95%) of thedissolved carbon dioxide. The so produced carbon dioxide is thenseparated in a separator (typically operating at atmospheric andsub-atmospheric pressure) from the lean solvent. It should be especiallyappreciated that the so generated carbon dioxide stream has a carbondioxide content of over 90%, and more typically of at least 95%. Thus,the so formed carbon dioxide stream is especially suited to be employedin enhanced oil recovery process.

In still further contemplated aspects of the inventive subject matter,the lean solvent from the separator is further let down in pressure viaJT valve and fed into a vacuum separator. Preferred vacuum separatorsoperate at a pressure of between about 1 to 10 psia, which may begenerated by a liquid seal vacuum pump. Residual carbon dioxide(typically with a purity of at least 95%) from the lean solvent isremoved in the vacuum separator and may also be employed in enhanced oilrecovery as depicted in FIGS. 3 and 4. The physical solvent is thenregenerated under the deep vacuum condition that may be assisted bystripping gas and recirculated to the absorber via a lean solvent pump.In particularly preferred configurations, the vacuum separator may use alean gas (e.g., a portion of the product gas) as a stripping gas toproduce an ultra lean solvent. However, in alternative configurations,various gases other than the product gas are also suitable, includinggases from other streams within the plant and even nitrogen or air. Itshould be further appreciated that the use of a vacuum separator incombination with a gas stripper in such configurations produces a verylean solvent capable of producing a treated gas with a CO₂ concentrationof typically less than 1000 ppmv.

Thus, contemplated configurations will provide pipeline quality gas athigh pressure and a carbon dioxide liquid stream, which can be used forenhanced oil recovery, wherein refrigeration is generated fromsuccessive depressurization of rich solvents. In especially preferredconfigurations, contemplated acid gas removal plants may operate withoutexternal refrigeration, and at higher pressure, such configurations willproduce refrigeration that can be used to condense carbon dioxide forfurther use in enhanced oil recovery. Besides providing refrigerant forremoving the heat of absorption from the absorber, the successivedepressurization will return the flash vapors containing methane andhydrocarbons to the absorber which are substantially fully recoveredduring the recycle process. Moreover, product gas from the absorber anddepressurized solvent at atmospheric pressure are employed to cool feedgas to the absorber maintaining the absorber bottom in a desirable lowtemperature range. It is therefore contemplated that the heat exchangeconfiguration produces an absorber temperature profile with either veryclose to isothermal or with a decreasing temperature profile, resultingin favorable physical properties that improve the column hydrodynamicperformance and absorption efficiency.

In particularly preferred configurations and where the feed gascomprises natural gas, it should be appreciated that the product gascomprises at least 90%, more typically at least 95%, and most typicallyat least 99% of the natural gas present in the feed gas. While notwishing to be bound be any theory or hypothesis, it is contemplated thatsuch relatively high natural gas recovery in the product gas is achievedby providing at least one, and more preferably threehydrocarbon-containing recycle streams back to the absorber, and/or byoperating the absorber under isothermal or a decreasing top-to-bottomthermal gradient. Suitable recycle gas compressors are all compressorsthat are capable of compressing the first and secondhydrocarbon-containing recycle gas streams to a pressure equal or aboutthe pressure of the cooled and dehydrated feed gas. Similarly, it iscontemplated that the lean solvent pump will provide solvent pressuresuitable for introduction of the lean solvent into the absorber.

Consequently, it is contemplated that configurations according to theinventive subject matter will significantly reduce overall energyconsumption and capital cost as compared to conventional carbon dioxideremoval processes at high carbon dioxide partial pressure using amine orother physical solvents or membranes. Moreover, contemplatedconfigurations and processes will generally not require an external heatsource or refrigeration, thereby further reducing energy consumption.Still further, enhanced oil recovery projects will frequently encounteran increase in carbon dioxide concentration in the feed gas, typicallyfrom 10% up to as high as 60%. Contemplated configurations and processescan accommodate these changes with essentially same solvent circulation.

A further advantage of contemplated configurations is that the processis generally a non-corrosive process due to operation at low temperatureand lack of water in the physical solvent. In contrast, conventionalamine units for carbon dioxide removal are generally more complex tooperate and maintain as such processes tend to be corrosive and oftenrequire antifoam and anti-corrosion injections during operation. Stillfurther, another advantage of contemplated physical solvent processes isthat, unlike amine processes, the solvent circulation rate is lesssensitive to increases in carbon dioxide partial pressure as the carbondioxide loading in the rich solvent merely increases with increasingcarbon dioxide concentration in the feed gas. In an amine unit design,the amine circulation rate would need to be increased linearly withincreasing carbon dioxide content.

Yet another advantage of contemplated physical solvent processes istheir simplicity and resistance to freezing compared to known aminetreating processes, thus requiring less supporting offsites and utilitysystems, such as steam boilers. For example, contemplated configurationsoperating a high carbon dioxide feed gas may not require any coolingduty as the flashing of carbon dioxide from the rich solvent willprovide the necessary cooling and regeneration. The inventors furthercontemplate that operation of a plant with vacuum regeneration canachieve a very low residual CO₂ content.

Thus, specific embodiments and applications for configurations andmethods for improved acid gas removal have been disclosed. It should beapparent, however, to those skilled in the art that many moremodifications besides those already described are possible withoutdeparting from the inventive concepts herein. The inventive subjectmatter, therefore, is not to be restricted except in the spirit of theappended contemplated claims. Moreover, in interpreting both thespecification and the contemplated claims, all terms should beinterpreted in the broadest possible manner consistent with the context.In particular, the terms “comprises” and “comprising” should beinterpreted as referring to elements, components, or steps in anon-exclusive manner, indicating that the referenced elements,components, or steps may be present, or utilized, or combined with otherelements, components, or steps that are not expressly referenced.

1. A method of removing an acid gas from a feed gas comprising: feedinginto an absorber the feed gas at a pressure of at least 400 psig,wherein the feed gas comprises at least 5 mol % carbon dioxide, whereinthe absorber is operated at an isothermal or decreasing top-to-bottomthermal gradient, wherein the absorber employs a physical solvent to atleast partially remove an acid gas from the feed gas to thereby producean absorber overhead product, a semi-rich solvent, and a rich solvent;cooling the semi-rich solvent using refrigeration content of at leastpartially expanded rich solvent; and cooling the feed gas usingrefrigeration content of the at least partially expanded rich solventand the absorber overhead product.
 2. The method of claim 1 wherein therich solvent is expanded in at least two steps, wherein expansion in onestep produces work, and wherein expansion in another step providesrefrigeration for the semi-rich solvent and optionally a carbon dioxideproduct.
 3. The method of claim 1 wherein the rich solvent is expandedin at least three steps, wherein expansion in the at least three stepsproduces at least three recycle streams, respectively, and wherein theat least three recycle streams are fed into the absorber.
 4. The methodof claim 3 wherein the at least three recycle streams are compressed toform a compressed recycle stream, and wherein further refrigeration isprovided by Joule-Thomson cooling of compressed recycle stream.
 5. Themethod of claim 1 wherein the feed gas is cooled by the at leastpartially expanded rich solvent.
 6. The method of claim 5 wherein thefeed gas is further cooled by the absorber overhead product.
 7. Themethod of claim 1 wherein at least part of the acid gas is removed fromthe physical solvent at a pressure of between about 1 psia to 10 psia.8. The method of claim 1 wherein the feed gas has a pressure betweenabout 400 psig to about 3000 psig, and wherein the feed gas is at leastpartially dehydrated.
 9. The method of claim 1 wherein the feed gas hasan acid gas content of between about 10 mol % to about 75 mol %.
 10. Themethod of claim 1 wherein the feed gas comprises natural gas.
 11. Themethod of claim 1 wherein the absorber is operated at a bottomtemperature of about −25° F. to about −45° F.
 12. The method of claim 1wherein the rich solvent is expanded to provide refrigeration for acarbon dioxide product.
 13. The method of claim 1 wherein the feed gashas a pressure of at least 1000 psig, and wherein at least a portion ofthe acid gas in the feed gas is removed from the feed gas using amembrane separator.
 14. A plant comprising: a gas source that isconfigured to provide natural gas comprising at least 5 mol % acid gasat a pressure of at least 400 psig; an absorber fluidly coupled to thegas source and configured to receive the natural gas and furtherconfigured to form a semi-rich solvent from a physical solvent; a coolerfluidly coupled to the absorber and configured to cool the semi-richsolvent and to provide the cooled semi-rich solvent back to the absorberat a temperature suitable to absorb at least another portion of the acidgas to thereby allow formation of a rich solvent in the absorber; firstand second expansion devices fluidly coupled to the absorber and furtherfluidly coupled to respective first and second heat exchangers, whereinthe first and second expansion devices and heat exchangers areconfigured to allow cooling of the the natural gas and the semi-richsolvent in the first and second heat exchangers, respectively, byexpansion of the rich solvent; and wherein the first and secondexpansion devices and heat exchangers are further configured to allowoperation of the absorber with an isothermal or decreasing top-to-bottomthermal gradient.
 15. The plant of claim 14 wherein the first and secondheat exchangers are configured such that cooling of the natural gas andthe semi-rich solvent provides an isothermal or decreasing top-to-bottomthermal gradient in the absorber.
 16. The plant of claim 14 wherein thefirst and second expansion devices are configured to allow formation ofat least one hydrocarbon containing recycle stream from the richsolvent, and further comprising a recycle compressor that is configuredto receive and compress the recycle stream to a pressure suitable forfeeding the recycle stream back to the absorber.
 17. The plant of claim16 further comprising a JT valve that is coupled to the recyclecompressor and configured to expand the compressed recycle stream tothereby provide refrigeration to the compressed recycle stream.
 18. Theplant of claim 16 further comprising a vacuum stripper that isconfigured to strip the solvent at a pressure of about 1 psia to about10 psia to thereby produce a lean solvent.
 19. The plant of claim 14wherein the gas source is configured to provide the natural gas at apressure of at least about 1000 psig, and further comprising a membraneseparator that is coupled to the gas source and configured to allowremoval of at least a portion of the acid gas in the natural gas. 20.The method of claim 1 further comprising a conduit fluidly coupled to arecycle gas cooler to allow withdrawal of a hydrocarbon liquid stream asa liquid product from the recycle gas.
 21. The plant of claim 14 furthercomprising a third expansion device that is configured to receive therich solvent and to allow for production of at least one of work andrefrigeration for a carbon dioxide product.